Process for the preparation of hydrogenated hydrocarbon compounds

ABSTRACT

A process for the dehydrogenation of a paraffinic hydrocarbon compound, such as an alkane or alkylaromatic hydrocarbon compound to produce an unsaturated hydrocarbon compound, such as an olefin or vinyl aromatic compound or mixture thereof, in which a dehydrogenation catalyst contacts gaseous reactant hydrocarbons in a reactor at dehydrogenation conditions.

This application is a continuation application of U.S. patentapplication Ser. No. 13/348,236, filed Jan. 11, 2012, which is acontinuation application of U.S. patent application Ser. No. 12/940,286,filed Nov. 5, 2010, which is a continuation of U.S. patent applicationSer. No. 10/586,024, filed Jul. 14, 2006, which is a 371 ofPCT/US2005/003772 filed Feb. 4, 2005, which claims the benefit of U.S.Provisional Application No. 60/543,006 filed 9 Feb. 2004.

BACKGROUND OF THE INVENTION

1) Field of the Invention

This invention generally relates to the field of hydrocarbon conversionand particularly to the dehydrogenation of paraffinic hydrocarbons toolefinic hydrocarbons, and/or lower alkylaromatic hydrocarbons to vinylaromatic hydrocarbons. In several preferred embodiments, the inventionrelates to the dehydrogenation of lower alkanes, for example ethane,isopropane, propane and butanes to their corresponding olefins, forexample ethylene, propylene and butylenes; and/or to the dehydrogenationof lower alkylaromatic hydrocarbon compounds, for example ethylbenzene,propylbenzene and methylethylbenzene to their corresponding vinylaromatic (that is “alkenylaromatic”) hydrocarbon compounds, for examplestyrene, cumene and alpha-methyl styrene, respectively. The inventionfurther includes an integrated process for making olefinic and vinylaromatic hydrocarbons including alkylation and dehydrogenation steps.

As required by 37 C.F.R. §1.71(g) and 37 C.F.R. §1.9(e), Applicantshereby assert that the claimed invention arose out of a joint researchagreement, as defined by 35 U.S.C. §103(c)(3), between SnamprogettiS.p.A. and The Dow Chemical Company. Said joint research agreement wasin effect on or before the effective filing date hereof, and the claimedinvention was made as a result of activities undertaken within the scopeof, and during the term of, the joint research agreement.

2) Description of Related Art

U.S. Pat. No. 6,031,143 and its corresponding EP 0 905 112 describe anintegrated process for producing styrene by feeding benzene and recycledethylene to an alkylation reactor to produce ethylbenzene, mixing thealkylation effluent with ethane and feeding the mixture to adehydrogenation reactor containing a catalyst capable ofcontemporaneously dehydrogenating ethane and ethylbenzene. The resultingproduct is separated to produce a stream of styrene and ethylene, withethylene being recycled to the alkylation reactor. The dehydrogenationreactor is preferably a fluidized bed reactor connected to a fluidizedbed regenerator from which the catalyst is circulated between theregenerator and the dehydrogenation reactor in countercurrent flow. Thatis, catalyst is introduced to the dehydrogenation reactor from the topand slowly descends to the bottom in countercurrent with the gas phasereactants which are rising through the reactor. During this descent, thecatalyst is deactivated. The deactivated catalyst is removed from thebottom of the dehydrogenation reactor and transported to the top of theregenerator where it descends to the bottom in countercurrent flow withhot air which is rising. During this descent, the carbonaceous residuepresent on the catalyst is burnt and the regenerated catalyst iscollected at the bottom of the regenerator where it is subsequentlycirculated back to the top of the dehydrogenation reactor.

WO 02/096844 describes an improvement to this process where thedehydrogenation catalyst is transported from the regenerator to thedehydrogenation reactor by way of a lower alkyl hydrocarbon carrier, forexample ethane. During transport, a portion of the carrier isdehydrogenated, (for example ethane converted to ethylene), and thecatalyst is cooled.

EP 1 255 719 (and corresponding co-pending US patent publication no. US2003/0028059, both filed by the assignee of the present application)describes a similar integrated process of preparing styrene usingbenzene and ethane as raw materials. However, the process includesadditional separation and recycling steps that are designed to improveefficiency. For example, the dehydrogenated effluent exiting thedehydrogenation reactor is separated into its aromatic and non-aromaticconstituents. The non-aromatic constituents, namely ethane, ethylene andhydrogen are recycled to an alkylation reactor were they are combinedwith benzene. The aromatic constituents are further separated, forexample styrene is recovered and ethylbenzene is recycled to thedehydrogenation reactor. The alkylation effluent is separated into itsconstituents with hydrogen being removed, and ethane and ethylbenzenebeing directed to the dehydrogenation reactor. The dehydrogenationreactor may have a variety of conventional designs including fixed,fluidized, and transport bed.

The described dehydrogenation processes are effective at integrating theproduction of styrene and ethylene using ethane and benzene as thestarting materials. Thus, these processes effectively de-coupled theproduction of styrene from the presence or proximity of a lighthydrocarbon steam cracker as a source for ethylene. However, thedehydrogenation processes described employ relatively long contact timesbetween the hydrocarbons and catalyst while at reaction temperature,resulting in thermal cracking, undesired side reactions and theformation of tars and other heavy products.

WO 02/096844 introduces the concept of a split “riser-type”dehydrogenation reactor operating in concurrent or “equicurrent” modewherein catalyst is carried upwards pneumatically through thedehydrogenation reactor by the gas phase reactants. The space velocity(GHSV) for such a reactor is greater than 500^(h-1). The catalyst isintroduced into the reactor with an alkyl hydrocarbon such as ethanewhereas the alkylaromatic compound, for example ethylbenzene, isintroduced at a suitable height along the riser after much of the alkylhydrocarbon has be dehydrogenated and the temperature of the catalysthas been reduced. While no specific examples or operating conditions areprovided, the use of such a riser reactor presumably leads to reducedcontact times between reactants and catalyst while in the reactor.

Dehydrogenation temperatures and residence times are typically optimizedto balance the reaction kinetics of both catalytic and gas-phase(thermal) reactions. The catalytic reaction produces a high selectivityto the desired products while the gas phase reaction produces manyundesired products and impurities. That is, while the catalytic reactionkinetics to the desired products increases exponentially withtemperature so does the gas phase reaction kinetics; therefore, theproper residence time and reaction temperature profile must be selectedto drive both the catalytic reaction to the desired conversion while notallowing the non-selective gas phase reactions to overwhelm the totalproduct selectivity. It would be useful to provide an apparatus andprocess which minimizes the time period in which reactants and catalystare in contact with one another while at reaction temperature. This isparticularly the case when utilizing highly reactive catalyst which canquickly deactivate.

While not directed toward a “dehydrogenation process” as described inthe aforementioned references, WO 03/050065 describes an integratedprocess for making styrene where benzene and “recycled” ethylene arecombined in an alkylation unit with the resulting product stream ofethylbenzene being combined with ethane. Unlike the previously describedreferences, this process utilizes an oxidative dehydrogenation(oxodehydrogenation) reaction. That is, the product stream from thealkylation unit is combined with ethane and oxygen and thencontemporaneously oxidatively dehydrogenated to provided ethylene andstyrene. The resulting ethylene is recycled to the alkylation unit. Theoxodehydrogenation reactor is described as a fluid-bed reactor operatingat a temperature range of from 300 to 550° C., a pressure range from 1to 30 bar, a gas hourly space velocity of 2000 to 6000^(h-1), with aresidence time of the catalyst in the fluid-bed zone of from 1 to 60seconds.

BRIEF SUMMARY OF THE INVENTION

The above described deficiencies of prior art can be overcome by thesubject invention which comprises contacting a gaseous stream ofhydrocarbon with a dehydrogenation catalyst at reaction temperature forrelatively short “contact times.” In a preferred embodiment, loweralkanes, for example ethane, propane and butanes are dehydrogenated totheir corresponding olefins, for example ethylene, propylene andbutylenes; and/or lower alkylaromatic hydrocarbon compounds, for exampleethylbenzene, propylbenzene and methylethylbenzene are dehydrogenated totheir corresponding vinyl aromatic hydrocarbon compounds, for examplestyrene, cumene and alpha-methyl styrene, respectively.

In another embodiment, the aforementioned dehydrogenation process iscombined with an alkylation step, as part of an integrated process. Manyadditional embodiments are also described.

BRIEF DESCRIPTION OF THE DRAWINGS

FIG. 1 shows a schematic block flow diagram of an embodiment of thepresent invention in which a riser reactor is employed in a singlehydrocarbon feed point which may be used for: 1) paraffinic hydrocarbon(for example ethane) feed only, 2) alkylaromatic hydrocarbon (forexample ethylbenzene) feed only, or 3) mixed feed (for example ethaneand ethylbenzene), including catalyst regeneration.

FIG. 2 shows a schematic block flow diagram of another embodiment of thepresent invention in which a riser reactor is employed with a multiplefeed point configuration, that is a split ethylbenzene and ethane feedconfiguration, including catalyst regeneration.

FIG. 3 shows a schematic block flow diagram of another embodiment of thepresent invention including multiple riser reactors with a catalystregeneration in a series configuration.

FIG. 4 shows a schematic block flow diagram of another embodiment of thepresent invention including multiple riser reactors with catalystregeneration in a parallel configuration.

FIG. 5 shows a schematic block follow diagram of another embodiment ofthe present invention similar to FIG. 4, but further including acatalyst recycle configuration.

DETAILED DESCRIPTION OF THE INVENTION

The present invention is directed toward the dehydrogenation of at leastone and preferably both of: 1) a paraffinic hydrocarbon compounds,preferably a lower alkane having from 2 to 6 carbon atoms but morepreferably less than 5 carbon atoms, for example ethane, propane,isopropane and butanes, to the corresponding olefin, namely, ethylene,propylene, and butylenes, respectively, and 2) an alkylaromatichydrocarbon compound, preferably a lower alkylaromatic hydrocarboncompound, such as for example, ethylbenzene, propylbenzene, isopropylbenzene, and methyl ethylbenzene, to the corresponding vinyl aromatichydrocarbon compound, (that is “alkenylaromatic”), namely, styrene,cumene or alpha-methyl styrene. Several embodiments of the presentinvention are described including both the simultaneous and separatedehydrogenation of lower alkanes and alkylaromatics. The invention isuseful to prepare styrene and ethylene from ethylbenzene and ethane,respectively. Likewise, cumene and propylene can be prepared frompropylbenzene and propane, respectively.

The dehydrogenation reaction in the present invention is conducted undera relatively short contact times in order to prevent undesirable sidereactions and product degradation. The term “average contact time” or“contact time” as used herein is intended to refer to the time in whichthe molar average of gaseous hydrocarbon molecules are in contact withcatalyst while at reaction temperature, regardless of whether thereactants are converted to desired products. The term “reactiontemperature” is intended to mean a temperature at which a significantamount of chemical reaction occurs, regardless of whether such reactionsare the desired dehydrogenation of reactants to their correspondingolefin and vinyl aromatic products. Said another way, the reactiontemperature is the temperature at which the hydrocarbons are no longerstable. The term “significant amount” in intended to mean a detectableamount having in an economic impact on the process. In most embodimentsof the invention, the reaction temperature is greater than about 500 andpreferably 550° C. The average contact time needs to be sufficientlylong to dehydrogenate acceptable amounts of hydrocarbon reactants butnot so long as to result in unacceptable amounts of by-products. Whilethe required contact time is related to the specific reactants,catalysts and reaction temperatures, in preferred embodiments of theinvention the contact time within the dehydrogenation reactor is lessthan 60 seconds, preferably from about 0.5 to about 10 seconds, morepreferably from about 1 to about 8 seconds, and still more preferablyfrom about 1 to about 4 seconds.

Due to the active nature of the preferred catalyst, the averageresidence time of the catalyst within the dehydrogenation reactor ispreferably less than about 60 seconds, preferably from about 0.5 toabout 40 seconds, more preferably about 1.0 to about 12.0 seconds, andstill more preferably from about 1.0 to about 10 seconds.

At such short catalyst residence times and average contact times in thedehydrogenation reactor, the temperature of the reaction mixture, whichmay be supplied in major part by the hot fresh or regenerated catalyst,is preferably from about 500 to about 800° C. With respect to loweralkanes, the reaction mixture is preferably from about 600 to about 750°C., and with respect to alkylaromatics from about 550 to 700° C. butmore preferably from about 570 to about 660° C. In general, the highesttemperature in the reactor will be found at its lower end and asreaction proceeds and the catalyst and reaction mixture ascends, thetemperature will decrease toward the upper end of the reactor.

The applicable operating pressure of the dehydrogenation reactor isquite broad, that is from about 3.7 to about 64.7 psia. The pressure atwhich the reaction proceeds is typically from about 14.7 to about 64.7psia, and preferably from about 14.7 to about 44.7 psia. However, inseveral preferred embodiments of the invention, the operating pressureof the dehydrogenation reactor may be below atmospheric, that is fromabout 3.7 to 14.7 psia, more preferably about 6.0 to about 14.7 psia.

The gas hourly space velocity (GHSV) for the present process has beenfound to range from about 1,000 to about 150,000 normal cubic meters/hrof hydrocarbon feed per cubic meter of catalyst at bulk density. Thesuperficial gas velocity of about 5 to about 80 ft/sec, preferably about15 to about 70 ft/sec. The catalyst flux is preferably about 10 to about120 lbs/ft²-sec with a catalyst to feed ratio of about 5 to about 100 ona weight to weight basis. The catalyst is preferably pneumatically movedthrough the reaction system by a carrier fluid, which is preferablyeither an inert diluent fluid or one of the reactants in gaseous form.Alternatively, the catalyst may be transported through the reactor undersub atmospheric pressure without diluent. Examples of inert diluentcarrier gases are nitrogen, volatile hydrocarbons for example methane,and other carriers which do not interfere with the reaction, steam,carbon dioxide, argon and the like. The paraffinic hydrocarbon compoundsuseful as reactants in the process of the present invention are alsopreferred carrier fluids and, most preferred are ethane, propane, andbutane. Steam is preferably not used in the present invention. Theamount of carrier gas required is only that amount necessary to maintainthe catalyst particles in fluidized state and transport the catalystfrom the regenerator to the reactor. Preferably, the amount of carriergas employed can range from about 0 to about 0.2 kg gas/kg catalyst.Injection points for carrier gas, especially reactant feed materialcarrier gas can be made at multiple points along the fresh orregenerated catalyst transfer line connecting the regenerator with thelower end of the riser reactor. The carrier gas will exit the reactorwith the product gas or through the vent stream of the regenerator. Inthe case where the carrier gas is also a reactant, a considerableportion of the carrier gas may be reacted and leave with the product gasstream from the reactor.

The short contact time required by the present invention can beaccomplished by way of a number of known reactor designs including fastfluidized, riser and downer reactors. Riser reactors are well known andcommonly employed in conversion of certain petroleum fractions intogasoline in fluidized bed catalytic cracking (FCC) processes. See forexample U.S. Pat. No. 3,888,762 which describes a short-timedilute-phase riser reactor designed for contact times of about 10seconds, and which further includes catalyst regeneration and recycleconfigurations—incorporated herein by reference. See also: USPublication No. 2004/0082824; WO 2001/85872 and WO 2004/029178. In anFCC process, a solid particulate catalyst, usually an acidic clay,silica-alumina or synthetic or natural zeolite type of catalyst, isintroduced with a carrier fluid to the lower end of a long, cylindricalor tubular reaction vessel together with a petroleum fraction atelevated temperature and moderate pressure. The cracking process occursin the petroleum as the liquid petroleum is vaporized by the hotcatalyst and both rise in the reactor cylinder. At the top of the riserreactor, the catalyst and hydrocarbon product are separated and thegasoline product stream exits via a vent pipe for separation and furtherprocessing into gasoline and heating oil fractions. The catalyst settlesin an annular space between the outside wall of the riser tube and aninner wall of the reactor housing through which a stripper gas contactsthe catalyst, at a rate which does not prevent settling of the catalyst,and strips off additional petroleum product from the catalyst surface.The catalyst is then sent to a regenerator/reactivator in which thecatalyst is contacted with a regeneration fluid, usually anoxygen-containing gas for combustion of any remaining hydrocarbons,heavy residuals or tars, and the regenerated catalyst is sent back tothe lower end of the riser reactor to contact additional petroleum forcracking. Spent catalyst may also be directly recycled to the lower endof the reactor without regeneration.

In a similar manner, in a preferred embodiment of the present inventionthe alkylaromatic hydrocarbon compound and/or the paraffinic hydrocarboncompound are introduced to the lower end of a reactor and contacted bythe hot fresh or regenerated catalyst which is pneumatically moved by acarrier gas. As the hydrocarbon compound(s) rise in the cylindricalreactor with the catalyst, the dehydrogenation reaction takes place andat the top or upper end of the riser, the vinyl aromatic hydrocarboncompound and/or lower olefin is separated from the catalyst. The riserreactor can be constructed from conventional materials used in FCC orpetrochemical processing and is conveniently a steel vessel using analloy sufficient for containing the hydrocarbon materials of thereaction, considering the temperature, pressure and flow rates employedand may be refractory lined. The dimensions of the riser reactor aredependent on the process design of a processing facility, including theproposed capacity or throughput, gas hourly space velocity (GHSV),temperature, pressure, catalyst efficiency and unit ratios of feedconverted to products at a desired selectivity.

The separation of gaseous hydrocarbon and catalyst is convenientlyaccomplished by means of a centrifugal impingement separator, such as acyclone separator, but the separation can by done by any conventionalmeans for solid-gas separations, including filtration and liquidsuspension. It is important to minimize the average contact time betweenthe catalyst and hydrocarbon once they exit the dehydrogenation reactor.This is preferably accomplished by at least one of two means; physicalseparation of catalyst from hydrocarbon, and cooling the catalyst and/orhydrocarbon to a temperature below the reaction temperature ofhydrocarbon present. The average contact time of the catalyst andhydrocarbon at reaction temperature in the separation device istypically less than 60 seconds, preferably less than about 10 seconds,and more preferably less than about 5 seconds, and still more preferablyless than about 3 seconds. The separation device may be a conventionalsolid-gas impingement separator, such as cyclone separators commonlyused in FCC applications. Preferred cyclone separators include twostaged or “coupled” designs including both positive and negativepressure designs. Further examples are provided in U.S. Pat. Nos.4,502,947; 4,985,136 and 5,248,411. Once separated, the catalyst iseither recycled to the dehydrogenation reactor or transferred to aregenerator.

In addition to separating the catalyst and hydrocarbon, the separationdevice may include a heat exchanger and/or quenching unit for deliveringa fluid to cool the catalyst and/or hydrocarbons to a temperature belowthe reaction temperature. Such fluid may be delivered via a conventionalquenching design including pressurized nozzles for delivering quenchingfluid, for example liquid styrene, water, and the like. Such quenchingtechnology is available from Stone & Webster and BP Amoco.

The average contact time between the catalysts and hydrocarbons while atreaction temperature through the entire dehydrogenation reactor andseparation device is preferably less than 60 seconds, more preferablyless than about 20 seconds, and still more preferably less than about 10seconds, and event more preferably less than about 7 seconds.

Once separated, the gaseous hydrocarbon is further separated, that isaromatics and non-aromatics, etc., which may be part of an integratedprocess as described in U.S. Pat. No. 6,031,143; WO 02/096844; and US2003/0028059. The spent catalyst may then optionally be sent to astripper, and then either to a regenerator or recycle loop, after whichthe catalyst is returned to the dehydrogenation reactor. Duringregeneration the catalyst is contacted with a regeneration fluid,usually an oxygen-containing gas and optionally a fuel source such asmethane or natural gas where remaining hydrocarbons, coke, heavyresidues, tar, etc. are removed from the catalyst, and the resultingregenerated catalyst is cycled back to the dehydrogenation reactor. Aportion of the spent catalyst may be cycled back to the dehydrogenationreactor without regeneration via a recycle loop. Recycled spent catalystmay be combined with regenerated catalyst as a means of controllingtemperature and catalyst activity within the dehydrogenation reactor.The combination of recycled and regenerated catalyst may be optimizedbased upon feedback from the output of the dehydrogenation reactor. Anexample of a means for controlling this combination is described in WO03/083014, incorporated herein by reference. Examples of bothregeneration and recycle configurations are provided in U.S. Pat. No.3,888,762 and US 2003/0196933, which are also incorporated herein byreference.

Preferred catalysts for use in the present invention are very active andare capable of dehydrogenating paraffin and alkylaromatic hydrocarbonsin less than a few seconds at ideal reaction temperatures. Preferredcatalyst include solid particulate type which are capable offluidization and, preferably, a those which exhibit Geldart Aproperties, as known in the industry. Gallium-based catalyst describedin U.S. Pat. No. 6,031,143 and WO 2002/096844 and are particularlypreferred in the present process and are incorporated herein byreference. One class of preferred catalyst for the dehydrogenationreaction is based on gallium and platinum supported on alumina in thedelta or theta phase, or in a mixture of delta plus theta phases, ortheta plus alpha phases, or delta plus theta plus alpha phases, modifiedwith silica, and having a surface area preferably less than about 100m²/g, as determined by the BET method known to those skilled in thefield. More preferably, the catalyst comprises:

-   -   i) from 0.1 to 34 percent by weight, preferably 0.2 to 3.8        percent by weight of gallium oxide (Ga₂O₃);    -   ii) from 1 to 200 parts per million (ppm), preferably 100 to 150        ppm by weight of platinum;    -   iii) from 0.05 to 5 percent be weight, preferably 0.1 to 1        percent by weight of an alkaline and/or earth-alkaline such as        potassium;    -   iv) from 0.08 to 3 percent by weight silica;    -   v) the balance to 100 percent being alumina.        Similar gallium-based catalyst are described in WO 2003/053567        which further includes manganese; and US 2004/02242945 which        further includes zinc, and EP-B1-0,637,578. The description of        the catalyst from these documents is expressly incorporated        herein by reference.

Another suitable catalyst for the dehydrogenation reaction is based onchromium and comprises:

-   -   i) from 6 to 30 percent, preferably, from 13 to 25 percent, by        weight of chromium oxide (Cr₂O₃);    -   ii) from 0.1 to 3.5 percent, most preferably, from 0.2 to 2.8        percent, by weight stannous oxide (SnO);    -   iii) from 0.4 to 3 percent, most preferably, from 0.5 to 2.5        percent, by weight of an alkaline oxide, for example, potassium        oxide;    -   iv) from 0.08 to 3 percent by weight silica;    -   v) the balance to 100 percent being alumina in the delta or        theta phase, or a mixture of delta plus theta phases, or theta        plus alpha phases, or delta plus theta plus alpha phases.

The catalysts mentioned hereinabove can be used as such or diluted withan inert material, for example, alpha-alumina, possibly modified withoxides of alkaline metals and/or silica, at a concentration of the inertproduct of between 0 and 50 percent by weight.

Details on the preparation of the aforementioned catalysts and theirmore preferred species can be found in U.S. Pat. No. 6,031,143 andEP-B1-0,637,578. Typically, the process of preparing the aforementioneddehydrogenation catalysts comprises dispersing precursors of thecatalytic metals, for example, solutions of soluble salts of thecatalytic metals onto the carrier consisting of alumina or silica. Anexample of dispersion can comprise impregnation of the carrier with oneor more solutions containing the precursors of gallium and platinum, orwith one or more solutions of the precursors of chromium and tin,followed by drying and calcination. An alternative method comprises ionadsorption followed by the separation of the liquid portion of theadsorption solution, drying, and activation of the resultant solid. Asanother alternative, the carrier can be treated with volatile species ofthe desired metals. In the case of added alkaline or alkaline earthmetals, the addition procedure comprises co-impregnation of the alkalineor alkaline earth metal with the primary catalytic metals (that is, Gaand Pt, or Cr and Sn), or alternatively, addition of the alkali oralkaline earth metal to the carrier prior to dispersion of the primarycatalytic metals, and thereafter, possible calcination of the solid.

Other suitable dehydrogenation catalysts, based on iron oxide, aredisclosed in EP 1 216 219. These catalyst comprise:

-   -   (i) from 1 to 60 percent, preferably from 1 to 20 percent, by        weight iron oxide;    -   (ii) from 0.1 to 20 percent, preferably from 0.5 to 10 percent,        by weight of at least one alkaline or alkaline earth metal        oxide, more preferably, potassium oxide;    -   (iii) from 0 to 15 percent, preferably, from 0.1 to 7 percent,        by weight of at least one rare earth oxide, preferably, selected        from the group consisting of cerium oxide, lanthanum oxide,        praseodymium oxide, and mixtures thereof;    -   (iv) the complement to 100 percent being a carrier consisting of        a microspheroidal alumina with a diameter selected from those in        delta or theta phase, or in a mixture of theta plus alpha        phases, or in a mixture of delta plus theta plus alpha phases,        modified preferably with from 0.08 to 5.0 weight percent of        silica.        The carrier in the preferred iron oxide catalyst more preferably        has an average particle diameter and particle density such that        the final product can be classified as Group-A according to        Geldart (Gas Fluidization Technology, D. Geldart, John Wiley &        Sons) and a surface area of less than about 150 m²/g, as        measured by the BET method known to those skilled in the art.        The process of preparing the iron oxide catalyst is well known        and fully described in EP 1 216 219

Another applicable dehydrogenation catalyst consists of a mordenitezeolite, optionally, promoted with a metal selected from gallium, zinc,the platinum group metals, or a combination thereof, as described inU.S. Pat. No. 5,430,211 and incorporated herein by reference. Themordenite is preferably acid extracted and thereafter impregnated orion-exchanged with one or more metals selected from gallium, zinc, andthe platinum group metals, more preferably, gallium. In this catalyst,the total metal loading typically ranges from 0.1 to 20 weight percent,based on the total weight of the catalyst.

As mentioned, the preferred catalyst for use with the present inventionare very active and are capable of completing the dehydrogenationreaction in a relatively short reaction time, for example in matter ofseconds. Consequently, if the catalyst is allowed to remain in contactwith the hydrocarbon mixture at reaction temperature for a longer periodthan necessary to complete the dehydrogenation reaction, undesirableby-products are formed from unreacted starting materials and/or thedesired products are degraded by a continued exposure to the catalyst atprocess conditions. The use of short contact times between thehydrocarbon and catalyst while at reaction temperature in thedehydrogenation reactor results in an unexpectedly beneficialconversion, selectivity and decrease in the amounts of by-productsformed. This unexpected effect is magnified by the use of short contacttimes between the hydrocarbon products and catalyst while at reactiontemperature in the separation device. Further, the use of a reactor withrelatively short contact or residence time decreases the amount ofcatalyst required for the process. A lower catalyst inventory providesoperating and capital advantages compared with prior art processes.

BRIEF DESCRIPTION OF THE SEVERAL VIEWS OF THE DRAWINGS

Several preferred embodiments of the invention are illustrated in theattached figures. Turning to FIG. 1, a tubular cylindrical riser reactor10 having a lower end 12 and an upper end 14 is connected at its lowerend 12 to a fresh or regenerated catalyst transfer line 16 and at itsupper end 14 to a product gas exit line 18. Spent or deactivatedcatalyst is removed from the product gas at upper end 14 by a separationdevice (not shown) which can be a conventional solid-gas impingementseparator, such as a cyclone separator as previously described, and thecatalyst is sent via spent catalyst transfer line 20 to regenerator 22which is a reaction vessel in which combustion air is blown into theregenerator 22 by means of air line 24. Supplemental fuel may be addedvia fuel line 62 to provide the heat of reaction and necessary sensibleheat, including the heat of vaporization in the case of liquid feed inthe riser reactor 10. The combustion products from the oxidation ofhydrocarbon on the catalyst are removed from the regenerator 22 by meansof vent gas line 28. Prior to being sent for disposal or additional heatrecovery, the vent gas may be filtered for removal of catalyst fines anddust by conventional equipment which is not shown. As a result of thecombustion and hydrocarbon removal the catalyst is regenerated andheated to a temperature sufficient to dehydrogenate the hydrocarbon feedmaterials and is removed from the regenerator 22 by means of regeneratedcatalyst exit line 30. Fluidization is maintained by injection of adiluent or carrier gas, for example nitrogen, by means of nitrogeninjection lines 26 and 32, and carrier gas injection lines 34, 36, and38, so that catalyst is introduced to the lower end 12 of riser reactor10 where it contacts ethane which is introduced via hydrocarbon feedline 40.

While FIG. 1 has been described with reference to the dehydrogenation ofethane, it will be appreciated that the present invention, along withthe embodiment of FIG. 1 is also applicable for the dehydrogenation ofother hydrocarbons, including lower alkanes such as propane and butane,and lower alkylaromatics, such as ethylbenzene, propylbenzene andmethylethylbenzene.

In operation, the embodiment shown in FIG. 1 proceeds by feedingregenerated catalyst at a temperature of from about 600 to about 800° C.from the regenerator 22 by means of regenerated catalyst exit line 30into fresh or regenerated catalyst transfer line 16 with the catalystbeing maintained in a fluid state of a Geldart A solid particulatematerial by means of fluidizing inert gas, such as nitrogen, fed throughnitrogen injection lines 26 and 32, and carrier gas, which may be inert(again, such as nitrogen) or a reactant gas, such as a paraffinichydrocarbon, such as for example, a lower alkane, preferably ethane,propane, or a butane, via carrier gas injection lines 34, 36, and 38.This catalyst and carrier gas mixture is introduced to the lower end 12of riser reactor 10 and contacts a hydrocarbon feed in liquid or gaseousform, preferably the latter, introduced by means of hydrocarbon feedline 40. The catalyst and hydrocarbon feed, for example, a lower alkane,such as ethane, propane or a butane, or an alkylaromatic hydrocarboncompound, or a mixture of both lower alkane and an alkylaromatichydrocarbon compound, contacts the catalyst and rises in the riserreactor 10 with the catalyst, feed (which by this time has beentransformed into a gas) and the carrier gas. As thecatalyst-feed-carrier gas mixture rises in the reactor, thedehydrogenation reaction occurs and the feed is converted into a lowerolefin and/or a vinyl aromatic compound, depending on the feed material.As the reaction mixture containing gas and catalyst arrives at the upperend 14 of riser reactor 10, the catalyst and gaseous reaction mixtureare separated by a solid-gas separation device, such as an impingementseparation device which may preferably be a cyclone gas-solid separator,which is conventional and not shown, but which is well known to those ofskill in the art of the FCC industry. The separated product gas is sentto recovery and purification and the catalyst is sent for regenerationand re-heating by means of spent or deactivated catalyst transfer line20. As the spent or deactivated catalyst is introduced into theregenerator 22, it contacts heated combustion air which is introduced byair line 24 and supplemental fuel introduced by fuel line 62, such thatthe hydrocarbon materials remaining on the surface of the catalyst areburned off and exit the regenerator via vent gas line 28. The combustionprocess also serves a second purpose and that is to heat the catalyst sothat the catalyst can function as a heat transfer agent or medium in theriser reactor 10. As used in this embodiment, the hydrocarbon feed 40can be a paraffinic hydrocarbon such as a lower alkane, an alkylaromatichydrocarbon compound, or a mixture of the two.

FIG. 2 illustrates another preferred, non-limiting embodiment which is avariant on the process of the present invention using a similar riserreactor 10 configuration as described with respect to FIG. 1. In thisembodiment the paraffinic hydrocarbon (for example ethane) is fed to theriser reactor 10 at or adjacent the lower end 12 by means of ethane feedline 44 and the lower alkylaromatic hydrocarbon compound (for exampleethylbenzene), is fed at a higher point in the riser reactor 10, forexample at ethylbenzene feed line 42. Thus, the type of reactionillustrated by the process of FIG. 2 is a “split feed” riser reactorprocess which produces styrene and by-products, such as ethylene whichcan be returned to an alkylation step to react with additional benzeneto produce more ethylbenzene as part of an integrated process.

FIG. 3 illustrates yet another preferred, non-limiting embodiment of theinvention. In this embodiment, a “dual riser” reactor configuration isillustrated in which the riser reactors 10 and 48 are connected inseries. As shown in FIG. 3, riser reactor 10 has lower end 12 and upperend 14. Connected to lower end 12 is fresh or regenerated catalyst line16 and the catalyst is maintained in fluidized state by injection ofcarrier gas via lines 34 and 36. Hydrocarbon feed material, such asethane, is introduced to the lower end 12 of riser reactor 10 by meansof hydrocarbon feed line 40. At this stage of the process, theconfiguration is much like that of FIG. 1; however, the product gas fromriser reactor 10 in FIG. 3 is fed to a separation and recovery section(not shown) by means of product gas exit line 18 from which a sideproduct gas line 46 leads to an alkylaromatic hydrocarbon compound feedline, such as ethylbenzene feed line 42. Alternatively, both sideproduct gas line 46, which carries primarily the lower olefin producedin riser reactor 10 in addition to by-products and carrier gas, can befed separately into a second riser reactor, such as at 48, having alower end 50 and an upper end 52. Also entering the lower end 50 ofsecond riser reactor 48 is a partially deactivated catalyst line 54which leads from the upper end 14 of riser reactor 10 to the lower end50 of second riser reactor 48. Carrier gas line 38 can be used tointroduce fluidizing carrier gas into partially deactivated catalystline 54 at one or multiple points along partially deactivated catalystline 54. As the ethylene and ethylbenzene rise in second riser reactor48 with the catalyst and carrier gas, the catalyst is at a lowertemperature than when initially introduced to the lower end 12 of riserreactor 10. The relatively lower temperature than in riser reactor 10permits satisfactory reaction rates for the alkylaromatic hydrocarboncompound and prevents over reaction to undesired by-products, thusdecreasing the yield, conversion and selectivity of the dehydrogenationreaction. The upper end 52 of second riser reactor 48 is connected tosecond product gas exit line 56 and can lead the vinyl aromatichydrocarbon compound, such as crude styrene monomer contained in theproduct gases, into the product gas separation and recovery section,which is conventional and not further described or identified herein.Prior to exit from second riser reactor 48, the reaction mixture must beseparated from the deactivated catalyst and this is done in a solid-gasseparation device, such as a cyclone separator, not shown. The separatedand deactivated catalyst is fed back to the regenerator 22 by means ofspent or deactivated catalyst transfer line 20 which in this embodimentleads from the upper end 52 of second riser reactor 48 to theregenerator 22 where the catalyst is regenerated, as previouslydescribed. In operation, the process is much like that described inrelation to the process illustrated in FIGS. 1 and 2, except that theproduct gas from the upper end 14 of riser reactor 10 is split and aportion is introduced into the lower end 50 of second riser reactor 48.Ethylbenzene is also introduced into the lower end 50 of second riserreactor 48 along with the partially deactivated catalyst via partiallydeactivated catalyst line 54 and the dehydrogenation of the ethylbenzeneproceeds at somewhat milder conditions in second riser reactor 48 thanin riser reactor 10. At the upper end 52 of second riser reactor 48, theproduct gases are separated from the catalyst in a solid gas separatordevice, such as a cyclone separator (which is conventional and notshown) and the product gases exit via second product gas exit line 56and the catalyst is sent back to regenerator 22 for regeneration andreheating via spent or deactivated catalyst transfer line 20.

In a still further preferred embodiment of this invention shown in FIG.4, the reactor/regenerator configuration is similar to that of FIG. 3,except that the second riser reactor 48 has its own catalyst feed andremoval transfer lines, namely second fresh or regenerated catalysttransfer line 58 and second spent or deactivated catalyst transfer line60 which feed active catalyst to second riser reactor 48 and removecatalyst from it and send the deactivated or spent catalyst back toregenerator 22. While shown as utilizing a common regenerator 22, itwill be appreciated that each reactor may include a separateregenerator.

In operation and as shown in FIG. 4, the catalyst from regenerator 22 isled by regenerated catalyst exit line 30 to either riser reactor 10 orsecond riser reactor 48 via fresh or regenerated catalyst transfer line16 or second fresh or regenerated catalyst transfer line 58,respectively. The feed to riser reactor 10 is ethane via hydrocarbonfeed line 40 and to second riser reactor 48 is ethylbenzene viaethylbenzene feed line 42. On contact with the catalyst in the riserreactors, the ethane and ethylbenzene are converted into ethylene andstyrene monomer, respectively, and the crude gaseous products areseparated from the catalyst in gas-solid separators, such as cycloneseparators (not shown) and sent to product gas separation and recoveryoperations (not shown) to produce ethylene for recycle to makeadditional ethylbenzene and styrene monomer, respectively. In a similarmanner and using propane or butane instead of ethane feed, the processof this invention would dehydrogenate the feed to propylene orbutylenes, respectively; or using isopropyl benzene or methyl ethylbenzene as feed material, the process of this invention woulddehydrogenate the feed to cumene or alpha-methyl styrene, respectively.

FIG. 5 illustrates yet another embodiment of the invention similar tothat shown in FIG. 4 but with the addition of a catalyst recycle loopcomprising a catalyst transfer line 64, carrier gas injector line 66 andflow valve 68. Spent catalyst is removed from the product gas at theupper end 52 of the dehydrogenation reactor 48 via a separation device(not shown) and is recycled back to the bottom end 50 of the reactor 48via catalyst transfer line 64. Fluidization of the spent catalyst ismaintained by the injection of a carrier gas, for example nitrogen bymeans of injection line in module 66. In addition to providing a carriergas, an oxygen-containing gas may be introduced in order to partiallyreactivate the catalyst, in which case module 66 would include a chamberfor reaction and removal of hydrocarbon residue. Flow of catalystthrough the recycle loop is controlled by one or more valves, forexample 68 which may be controlled remotely according to predeterminedperformance criteria including reactor 48 temperature, catalystactivity, etc. Recycled catalyst may be combined with regeneratedcatalyst prior to introduction in the bottom of reactor 48, or may beintroduced via separate entry points (not shown).

Additional configurations of the reactor(s), regenerator and recycleloop can be envisioned by one skilled in the art. For example, oneskilled in the art will appreciate that multiple reactors could bearranged to feed into a common separation device with shared or separatecatalyst regenerators and various recycle loops. The present inventionis desired to be limited only by the lawful scope of the appended claims

The present invention does not preferably include oxidativedehydrogenation, that is oxodehydrogenation. In fact, oxygenates canpoisen some types of catalyst; however, oxygen may used to regenerate orreactive catalyst during the regeneration process. Moreover, the presentinvention preferably does not utilize steam as is typically used inconvention styrene production process.

Another preferred embodiment of the invention utilizes the previouslydescribed dehydrogenation process as part of an integrated process formaking olefins and vinyl aromatics. More specifically, the previouslydescribed dehydrogenation, (along with regeneration and/or recycleprocesses) can be used to replace the dehydrogenation schemes describedin U.S. Pat. No. 6,031,143; WO 02/096844; and co-pending US2003/0028059. In such an integrated process, a paraffinic hydrocarbonsuch as a lower alkane, for example ethane, and benzene are the primaryraw materials. Ethylene, preferably “recycled” and benzene are feed to aconventional alkylation reactor as is well known in the art and asdescribed in the references mentioned above, hereby incorporated byreference. Alkylation of benzene with ethylene in typically conducted inthe presence of aluminum chloride or zeolites catalyst. Variationsinclude the use of dilute ethylene and a catalytic distillation approachwhere liquid phase alkylation and product separation take placesimultaneously. Specific examples include the “EBOne Process” availablefrom ABB Lummus/UOP, “EB Max Process” available from ExxonMobil/Badgerand similar alkylation technology available from CDTECH, a partnershipbetween ABB Lummus Global Inc. and Chemical Research and Licensing.

The alkylation affluent is recovered and optionally subject toseparation, that is separation of aromatics from non-aromatics, removalof hydrogen, etc. Alkylaromatic, for example ethylbenzene, andparaffinic hydrocarbon, for example ethane, are then dehydrogenation aspreviously described. The gaseous products of dehydrogenation arerecovered and separated, for example aromatics from non-aromatics, withvinyl aromatics, for example styrene being recovered, olefins, forexample ethylene (and possibly paraffinic hydrocarbons, for exampleethane) being recycled to the alkylation reactor, and alkylaromaticsbeing recycled to the dehydrogenation reactor.

The invention claimed is:
 1. A process for dehydrogenating a hydrocarbonselected from at least one of: i) paraffinic hydrocarbons selected fromethane, propane, and butane; and ii) alkylaromatic hydrocarbons selectedfrom ethylbenzene, propylbenzene and methylethylbenzene; comprisingcontacting a gaseous stream containing at least one of the hydrocarbonswith a dehydrogenation catalyst comprising gallium and platinum andcarried by an alumina or alumina silica support, at reaction temperatureand in concurrent rising flow, at a catalyst to gaseous stream ratio of5 to 100 on a weight to weight basis, to a dehydrogenation reactorwherein the average contact time between the hydrocarbon and catalystwithin the dehydrogenation reactor is from about 1 to about 8 seconds;and the temperature and pressure in the dehydrogenation reactor is fromabout 500 to about 800° C., and from about 14.7 to about 44.7 psia;transferring the hydrocarbon and catalyst from the dehydrogenationreactor to a two-staged solid-gas separation device; and transferringcatalyst from the solid-gas separation device to a regenerator where thecatalyst is contacted with combustion air and supplemental fuel.
 2. Theprocess of claim 1 wherein the dehydrogenation reactor is a riserreactor.
 3. The process of claim 1 wherein the catalyst comprises analkali or alkaline earth metal.